Process for methane conversion

ABSTRACT

A process for converting methane to higher hydrocarbon(s) including aromatic hydrocarbon(s) in a reaction zone comprises providing to a hydrocarbon feedstock containing methane and a catalytic particulate material to the reaction zone and contacting the catalytic particulate material and the hydrocarbon feedstock in a substantially countercurrent fashion in the reaction zone, while operating the reaction zone under reaction conditions sufficient to convert at least a portion of said methane to a first effluent having said higher hydrocarbon(s).

CROSS REFERENCE TO RELATED APPLICATIONS

This application claims the benefit of Provisional Application No.60/794,280 filed Apr. 21, 2006, the disclosures of which areincorporated by reference in their entireties.

FIELD

This disclosure relates to a process for methane conversion. Inparticular, this disclosure relates to a process for natural gasconversion.

BACKGROUND

Aromatic hydrocarbon(s), particularly benzene, toluene, ethylbenzene andxylenes, are important commodity chemicals in the petrochemicalindustry. Currently, aromatics are most frequently produced frompetroleum-based feedstocks by a variety of processes, includingcatalytic reforming and catalytic cracking. However, as the worldsupplies of petroleum feedstocks decrease, there is a growing need tofind alternative sources of aromatic hydrocarbon(s).

One possible alternative source of aromatic hydrocarbon(s) is methane,which is the major constituent of natural gas and biogas. World reservesof natural gas are constantly being upgraded and more natural gas iscurrently being discovered than oil. Because of the problems associatedwith transportation of large volumes of natural gas, most of the naturalgas produced along with oil, particularly at remote places, is flaredand wasted. Hence the conversion of alkanes contained in natural gasdirectly to higher hydrocarbon(s), such as aromatics, is an attractivemethod of upgrading natural gas, providing the attendant technicaldifficulties can be overcome.

A large majority of the processes currently proposed for convertingmethane to liquid hydrocarbon(s) involve initial conversion of themethane to synthesis gas, a blend of H₂ and CO. However, production ofsynthesis gas is capital and energy intensive and hence routes that donot require synthesis gas generation are preferred.

A number of alternative processes have been proposed for directlyconverting methane to higher hydrocarbon(s). One such process involvescatalytic oxidative coupling of methane to olefins followed by thecatalytic conversion of the olefins to liquid hydrocarbon(s), includingaromatic hydrocarbon(s). For example, U.S. Pat. No. 5,336,825 disclosesa two-step process for the oxidative conversion of methane to gasolinerange hydrocarbon(s) comprising aromatic hydrocarbon(s). In the firststep, methane is converted to ethylene and minor amounts of C₃ and C₄olefins in the presence of free oxygen using a rare earth metal promotedalkaline earth metal oxide catalyst at a temperature between 500° C. and1000° C. The ethylene and higher olefins formed in the first step arethen converted to gasoline range liquid hydrocarbon(s) over an acidicsolid catalyst containing a high silica pentasil zeolite.

However, oxidative coupling methods suffer from the problems that theyinvolve highly exothermic and potentially hazardous methane combustionreactions and they generate large quantities of environmentallysensitive carbon oxides.

A potentially attractive route for upgrading methane directly intohigher hydrocarbon(s), particularly ethylene, benzene and naphthalene,is dehydroaromatization or reductive coupling. This process typicallyinvolves contacting the methane with a catalyst comprising a metal, suchas rhenium, tungsten or molybdenum, supported on a zeolite, such asZSM-5, at high temperature, such as 600° C. to 1000° C. Frequently, thecatalytically active species of the metal is the zero valent elementalform or a carbide or oxycarbide.

For example, U.S. Pat. No. 4,727,206 discloses a process for producingliquids rich in aromatic hydrocarbon(s) by contacting methane at atemperature between 600° C. and 800° C. in the absence of oxygen with acatalyst composition comprising an aluminosilicate having a silica toalumina molar ratio of at least 5:1, said aluminosilicate being loadedwith (i) gallium or a compound thereof and (ii) a metal or a compoundthereof from Group VIIB of the Periodic Table.

In addition, U.S. Pat. No. 5,026,937 discloses a process for thearomatization of methane which comprises the steps of passing a feedstream, which comprises over 0.5 mole % hydrogen and 50 mole % methane,into a reaction zone having at least one bed of solid catalystcomprising ZSM-5, gallium and phosphorus-containing alumina atconversion conditions which include a temperature of 550° C. to 750° C.,a pressure less than 10 atmospheres absolute (1000 kPa-a) and a gashourly space velocity of 400 to 7,500 hr⁻¹.

Moreover, U.S. Pat. Nos. 6,239,057 and 6,426,442 disclose a process forproducing higher carbon number hydrocarbon(s), e.g., benzene, from lowcarbon number hydrocarbon(s), such as methane, by contacting the latterwith a catalyst comprising a porous support, such as ZSM-5, which hasdispersed thereon rhenium and a promoter metal such as iron, cobalt,vanadium, manganese, molybdenum, tungsten or a mixture thereof. Afterimpregnation of the support with the rhenium and promoter metal, thecatalyst is activated by treatment with hydrogen and/or methane at atemperature of about 100° C. to about 800° C. for a time of about 0.5hr. to about 100 hr. The addition of CO or CO₂ to the methane feed issaid to increase the yield of benzene and the stability of the catalyst.

WO 03/000826 and U.S. Patent Application Publication No 2003/0083535disclose a system and method for circulating catalyst between a reactorsystem and a regenerator system. A circulating catalyst system includesa reactor system, a regenerator system, and a distribution unit. Thereactor system and regenerator system are adapted to exchange catalyst.The reactor system preferably includes a fluidized bed riser reactor andthe regeneration system preferably includes a regeneration zone adaptedfor the contact of catalyst with a regeneration gas. The system andmethod are adapted so that more than one regeneration gas may contactcatalyst. The distribution unit is adapted to control the percentage ofcatalyst contacting each regeneration gas. Thus, the distribution unitis adapted to select the percentage so as to maintain the reactor systemand regeneration system under a heat balance regime. Heat is preferablytransferred from the regenerator system to the reactor system by anexchange of catalyst.

The successful application of reductive coupling to produce higherhydrocarbons, e.g., aromatic compounds, on a commercial scale requiresthe solution of a number of serious technical challenges. Examples ofthese technical challenges are:

(a) the process is endothermic which requires high energy input;

(b) the process is thermodynamically limited, which requires hightemperature operation to achieve high conversion;

(c) the process requires significant amounts of make-up heat tocompensate the energy requirement of the endothermic reaction and tomaintain the high temperature required for high conversion;

(d) the process requires effective heat transfer and effective contactof light hydrocarbon(s) with the catalyst to achieve high conversion ofmethane;

(e) the process generates coke and/or catalyst coking at hightemperature;

(f) the process may use feedstocks containing C₂+ hydrocarbons inaddition to methane, which feedstocks may increase coking of thecatalyst used in the process; and

(g) to reduce problems related to catalyst attrition, it is desirable tominimize the circulation rate and other mechanical stresses on thecatalyst.

Accordingly, there is a need to develop a process for converting methaneto higher hydrocarbon(s), which provides high efficiency for heattransfer, adequate hydrocarbon/catalyst contacting, improved processconditions to maximize selectivity to desired higher hydrocarbons, e.g.,aromatic compound(s), while minimizing coke formation, and minimizing ofrequired catalyst circulation rates.

SUMMARY

In one aspect, the present disclosure resides in a process forconverting methane to higher hydrocarbon(s) including aromatichydrocarbon(s) in a reaction zone, the process comprising:

(a) providing to said reaction zone a hydrocarbon feedstock containingmethane;

(b) providing to said reaction zone a catalytic particulate material;

(c) contacting said catalytic particulate material and said hydrocarbonfeedstock in a substantially countercurrent fashion;

(d) maintaining the hydrodynamics of said reaction zone in settling bedregime; and

(e) operating said reaction zone under reaction conditions sufficient toconvert at least a portion of said methane to a first effluent havingsaid higher hydrocarbon(s).

Additionally, the process may further comprise:

(f) removing at least a portion of said catalytic particulate materialfrom said reaction zone; and

(g) regenerating at least a portion of the removed catalytic particulatematerial under regenerating conditions.

Conveniently, the process further comprises recycling at least a portionof the regenerated catalytic particulate material to said reaction zone.

Alternatively, the process further comprises:

(f) removing at least a portion of said catalytic particulate materialfrom said reaction zone; and

(g) heating at least a portion of the removed catalytic particulatematerial to a temperature from about 800 to 1200° C.

Additionally, the process further comprises:

(h) recycling at least a portion of the heated catalytic particulatematerial to said reaction zones.

In a further aspect, this disclosure resides in a process for convertingmethane to higher hydrocarbon(s) including aromatic hydrocarbon(s) in areaction zone, the process comprising:

(a) providing to said reaction zone a hydrocarbon feedstock containingmethane;

(b) providing to said reaction zone a catalytic particulate material;

(c) contacting said catalytic particulate material and said hydrocarbonfeedstock in a substantially countercurrent fashion;

(d) maintaining the temperature profile of said reaction zone with aninverse temperature profile; and

(e) operating said reaction zone under reaction conditions sufficient toconvert at least a portion of said methane to a first effluentcomprising said higher hydrocarbon(s).

In yet a further aspect, this disclosure resides in a process forconverting methane to higher hydrocarbon(s) including aromatichydrocarbon(s) in a reaction zone, the process comprising:

(a) supplying said reaction zone with a hydrocarbon feedstock containingmethane and a catalytic particulate material in a substantiallycountercurrent fashion;

(b) operating said reaction zone in a fashion substantially free offluidization and under reaction conditions effective to convert at leasta portion of said methane to a first effluent having said higherhydrocarbon(s);

(c) removing at least a portion of said catalytic particulate materialfrom said reaction zone;

(d) regenerating at least a portion of the removed catalytic particulatematerial under regenerating conditions;

(e) heating at least a portion of the removed catalytic particulatematerial and/or at least a portion of the regenerated catalyticparticulate material to a temperature from about 800 to 1200° C.; and

(f) recycling at least a portion of the heated catalytic particulatematerial of step (f) to said reaction zone.

In still a further aspect, this disclosure resides in a process forconverting methane to higher hydrocarbon(s) including aromatichydrocarbon(s) in a reaction zone, the process comprising:

(a) supplying said reaction zone with a hydrocarbon feedstock containingmethane and a catalytic particulate material in a substantiallycountercurrent fashion;

(b) operating said reaction zone under reaction conditions effective toconvert at least a portion of said methane to a first effluent havingsaid higher hydrocarbon(s);

(c) removing at least a portion of said catalytic particulate materialfrom said reaction zone;

(d) heating at least a portion of the removed catalytic particulatematerial to a temperature from about 800 to 1200° C.; and

(e) regenerating at least a portion of the removed catalytic particulatematerial under regenerating conditions; and

(f) recycling at least a portion of the heated catalytic particulatematerial of step (d) and/or at least a portion of said regeneratedcatalyst of step (e) to said reaction zone.

Additionally, this disclosure resides in a process for manufacturingaromatic hydrocarbon(s) from methane in a reaction zone, the processcomprises:

(a) providing to said reaction zone a hydrocarbon feedstock containingmethane;

(b) providing to said reaction zone a catalytic particulate material;

(c) contacting said catalytic particulate material and said hydrocarbonfeedstock in a substantially countercurrent fashion;

(d) maintaining the temperature profile of said reaction zone as ainverse temperature profile; and

(e) operating said reaction zone under reaction conditions sufficient toconvert at least a portion of said methane to a first effluent havingsaid aromatic hydrocarbon(s); and

(f) recovering said aromatic hydrocarbon(s) from said effluent.

In some additional embodiments, the process comprises at least oneadditional reaction zone.

In some aspects of above mentioned embodiments, the process comprisesseparating unreacted methane from said higher hydrocarbon(s) andrecycling said unreacted methane to said reaction zone.

In additional aspects of above mentioned embodiments, the first effluenthas hydrogen and the process further comprises reacting at least part ofsaid hydrogen from said first effluent with oxygen-containing specie(s)to produce a second effluent having a reduced hydrogen content comparedwith said first effluent. Optionally, the second effluent is recycledback to step (a).

In one aspect of above mentioned embodiments, the reaction zone isoperated at a gas velocity of less than 0.99 times of the minimumfluidizing velocity.

In another aspect of above mentioned embodiments, (a) further comprisesa step of providing said reaction zone a non-catalytic particulatematerial.

Conveniently, the mass ratio of the total flowrate of said particulatematerial (catalytic particulate material plus any non-catalyticparticulate material) to the flowrate of said hydrocarbon feedstock isfrom about 1:1 to about 40:1, preferably from about 5:1 to 20:1

In yet another aspect of above mentioned embodiments, the reaction zoneis located in a cold wall reactor.

In another aspect of above mentioned embodiments, the reactionconditions are non-oxidizing conditions.

In one aspect of above mentioned embodiments, the reaction conditionsinclude a temperature of about 400° C. to about 1200° C., a pressure ofabout 1 kPa-a to about 1000 kPa-a, and a weight hourly space velocity ofabout 0.01 hr⁻¹ to about 1000 hr⁻¹.

In one embodiment of this disclosure, the reaction conditions aresufficient to convert at least 5 wt. % of said methane to said aromatichydrocarbon(s).

In another embodiment of this disclosure, the aromatic hydrocarbon(s)includes benzene.

In one embodiment of this disclosure, the catalytic particulate materialis a dehydrocyclization catalyst. In some embodiments, the catalyticparticulate material comprises a metal or compound thereof on aninorganic support. In other embodiments of this disclosure, the metal orcompound thereof comprises at least one of calcium, magnesium, barium,yttrium, lanthanum, scandium, cerium, titanium, zirconium, hafnium,vanadium, niobium, tantalum, chromium, molybdenum, tungsten, manganese,rhenium, iron, ruthenium, cobalt, rhodium, iridium, nickel, palladium,copper, silver, gold, zinc, aluminum, gallium, germanium, indium, tin,lead, bismuth or a transuranium element. In one aspect, the inorganicsupport comprises a microporous or mesoporous material. In someembodiments, the catalytic particulate material comprises at least oneof molybdenum, tungsten, rhenium, a molybdenum compound, a tungstencompound, a zinc compound, and a rhenium compound on ZSM-5, silica or analuminum oxide. In other embodiments, at least 90 wt. % said catalyticparticulate material has particle size from about 0.1 mm to about 100mm, optionally at least 90 wt. % said catalytic particulate material hasparticle size from about 1 to about 5 mm. In other embodiments of thisdisclosure, at least 90 wt. % said catalytic particulate material hasparticle size from about 2 to about 4 mm.

In some embodiments of this disclosure, said catalytic particulatematerial enters said reaction zone at a temperature of about 800° C. toabout 1200° C. and exits said reaction zone at a temperature of about500° C. to about 800° C. Conveniently, the temperature difference ofsaid catalytic particulate material across said reaction zone is atleast 100° C.

In some embodiments of this disclosure, the hydrocarbon feedstockfurther comprises at least one of CO₂, CO, H₂, H₂O, or C₂+hydrocarbon(s).

In some embodiments of this disclosure, the regeneration conditionscomprise a temperature from about 400° C. to about 750° C., and aregeneration gas having oxygen. Optionally, the regeneration gas furthercontains carbon dioxide and/or nitrogen such that the oxygenconcentration of said regeneration gas is from about 2 wt. % to about 10wt. %.

In some embodiments of this disclosure, at least a portion of saidaromatic hydrocarbon(s) is further contacted with a feed containinghydrogen to produce a product having saturates and/or single ringaromatic hydrocarbon(s).

In other embodiments of this disclosure, at least a portion of saidaromatic hydrocarbon(s) is further contacted with a feed containing analkylating agent to produce a product having xylene(s), ethylbenzene,cumene, or toluene.

These and other facets of the present disclosure shall become apparentfrom the following detailed description, figure, and appended claims.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a diagram of a process for converting methane to higherhydrocarbon(s) according to one embodiment of this disclosure.

DETAILED DESCRIPTION OF THE EMBODIMENTS

Definitions

All patents, patent applications, test procedures, priority documents,articles, publications, manuals, and other documents cited herein arefully incorporated by reference to the extent such disclosure is notinconsistent with the present disclosure and for all jurisdictions inwhich such incorporation is permitted.

When numerical lower limits and numerical upper limits are listedherein, ranges from any lower limit to any upper limit are contemplated.

As used in this specification, the term “framework type” is used in thesense described in the “Atlas of Zeolite Framework Types,” 2001.

As used herein the term “higher hydrocarbon(s)” means hydrocarbon(s)having more than one carbon atom per molecule, oxygenate having at leastone carbon atom per molecule, e.g., ethane, ethylene, propane,propylene, benzene, toluene, xylenes, naphthalene, and/or methylnaphthalene.

As used herein the term “aromatic hydrocarbon(s)” means molecule(s)containing one or more aromatic ring(s). Examples of aromatichydrocarbons are benzene, toluene, ethylbenzene, xylenes (para-xylene,meta-xylene, and ortho-xylene), naphthalene, and methyl naphthalene.

As used herein the term “moving bed” means a zone or vessel whereinparticulates contact with gas flows such that the superficial gasvelocity (U) is below the velocity required for dilute-phase pneumaticconveying of solid particles in order to maintain a solids bed with voidfraction below about 95 vol. % (U_(95%)), optionally below about 85 vol.% (U_(85%)). A moving-bed reactor may operate under several flow regimesincluding:

(a) settling- or moving packed-bed regime, wherein the superficial gasvelocity is less than the minimum fluidization velocity (U_(mf)),U<U_(mf);

(b) bubbling regime, wherein the superficial gas velocity is more thanthe minimum fluidization velocity (U_(mf)) and less than the minimumbubbling velocity (U_(mb)), U_(mf)<U<U_(mb);

(c) slugging regime, wherein the superficial gas velocity is more thanthe minimum bubbling velocity (U_(mb)) and less than the minimumbubbling velocity (U_(c)), U_(mb)<U<U_(c);

(d) transition to and turbulent fluidization regime, wherein thesuperficial gas velocity is more than the minimum velocity (U_(c)) andless than the minimum transport velocity (U_(tr)), U_(c)<U<U_(tr); and

(e) fast-fluidization regime, wherein the superficial gas velocity ismore than the minimum transport velocity (U_(tr)), U>U_(tr).

These different flow regimes have been described in, for example,Chapter 3 of “Fluidization Engineering,” D. Kunii and O. Levenspiel,2^(nd) Edition, Butterworth-Heinemann, Boston, 1991 and Chapter 6 of“Chemical Process Equipment,” S. M. Walas, Butterworth-Heinemann,Boston, 1990, the entirety of which are incorporated by reference.

As used herein the term “settling bed” means a zone or vessel whereinparticulates contact with gas flows such that the superficial gasvelocity (U) is below the minimum velocity required to fluidize thesolid particles, the minimum fluidization velocity (U_(mf)), U<U_(mf),in at least a portion of the reaction zone, and/or operating at avelocity higher than the minimum fluidization velocity while maintaininga gradient in gas and/or solid property (such as, temperature, gas orsolid composition, etc.) axially up the reactor bed by using reactorinternals to minimize gas-solid back-mixing. Description of the minimumfluidization velocity is given in, for example Chapter 3 of“Fluidization Engineering,” D. Kunii and O. Levenspiel, 2^(nd) Edition,Butterworth-Heinemann, Boston, 1991 and Chapter 6 of “Chemical ProcessEquipment,” S. M. Walas, Butterworth-Heinemann, Boston, 1990, theentirety of which are incorporated by reference.

As used herein the term “fluidizing bed” means a zone or vessel whereinparticulates contact with gas flows such that the superficial gasvelocity (U) is sufficient to fluidize solid particles (i.e., above theminimum fluidization velocity U_(mf)) and is below the velocity requiredfor dilute-phase pneumatic conveying of solid particles in order tomaintain a solids bed with void fraction below about 95%.

As used herein the term “cascade moving beds” means a series arrangementof individual moving beds as the particulates or gas cascades from onemoving bed to another.

As used herein the term “cascade fluidizing beds” means a seriesarrangement of individual fluidizing beds such that there can be agradient in gas and/or solid property (such as, temperature, gas orparticulates composition, pressure etc.) as the particulates or gascascades from one fluidizing bed to another.

As used herein the term “riser” means a zone or vessel (such as,vertical cylindrical pipe) used for net upwards transport ofparticulates in fast-fluidization or pneumatic conveying fluidizationregimes. Fast fluidization and pneumatic conveying fluidization regimesare characterized by superficial gas velocities (U) greater than theminimum transport velocity (U_(tr)). Fast fluidization and pneumaticconveying fluidization regimes are described in, for example, Chapter 3of “Fluidization Engineering,” D. Kunii and O. Levenspiel, 2nd Edition,Butterworth-Heinemann, Boston, 1991 and Chapter 6 of “Chemical ProcessEquipment,” S. M. Walas, Butterworth-Heinemann, Boston, 1990, theentirety of which are incorporated by reference.

As used herein the term “catalytic particulate material” means arefractory material that causes an increase in the reaction rate of thefeed to the desired products at the process conditions. The catalyticparticulates material may form particulates without binder or be boundby an inorganic binder such as clay, silica, alumina, zirconia, or othermetal oxide, to maintain the physical integrity of the particles.Preferably the particles are of a substantially spherical shape. Theparticles may contain additional components to provide useful functionsby adjusting the thermal conductivity, the density, the heat capacity,and/or the attrition resistance of the catalytic particulate material tosecure a desired catalyst performance.

As used herein the term “non-catalytic particulate material” means aparticulate material which is not a catalytic particulate material. Thenon-catalytic particulate material may comprise a refractory inorganicmaterial that does not cause an increase in the reaction rate of thefeed to the desired products at the process conditions.

As used herein the term “cold wall reactor” or “cold wall vessel” meansa reactor or a vessel constructed with one or more layers of insulatingmaterial between the catalyst of the process and the metallic shellwhich acts as a pressure containment for the process; whereby atemperature gradient occurs across the insulating material so that themetallic shell is at a substantially lower temperature, such as morethan 50° C., preferably more than 100° C., more preferably more than300° C., and optionally more than 600° C., than the temperature of thecontained material.

As used herein the term “non-oxidizing conditions” means conditionswherein oxidizing agents (such as, O₂, NO_(x) and metal oxides which canrelease oxygen to oxidize methane to CO_(x)) are present at less than5%, such as at less then 1%, and typically at less than 0.1%, of theamount required for stoichiometric oxidation of the methane in the feed.

By “supplemental source of fuel” is meant that the source fuel isphysically separate from the catalyst and hence is not, for example,coke generated on the catalyst as a by-product of the dehydrocyclizationreaction.

As used herein the term “carburizing gas” means any gas that, under theconditions in a catalyst treatment zone, can convert at least a portionof the catalytic metal(s) in the catalytic particulate material from anoxidized state to an elemental form, to a carbidic species, or to a lessoxidized form. The carburizing gas can also partly coke active catalyticsites of the catalyst. Such active catalytic sites comprise catalyticmetal and/or other active sites capable of catalyzing the desiredreaction. The carburizing gas may also deposit a quantity of carbonand/or hydrocarbon species on the catalyst particulate material. Suchcarbon and/or hydrocarbon species may be intermediates to the formationof the desired higher hydrocarbon(s), e.g., aromatic compound(s). Thecarburizing gas comprises hydrocarbons, H₂, CO, CO₂, and any combinationthereof, such that the carburizing gas contains a source of both theelement carbon and the element hydrogen.

Introduction

The present disclosure provides a process for producing higherhydrocarbon(s), e.g., aromatic compound(s), by contacting a feedstockcontaining methane, typically together with H₂, CO and/or CO₂, with aparticulate dehydrocyclization catalyst in a reaction zone, typicallyoperated in a settling bed regime, under conditions effective to convertthe methane to higher hydrocarbon(s) and hydrogen, wherein saidconditions include substantially countercurrent flow of the feedstockand the particulate dehydrocyclization catalyst in the reaction zone.

Generally, the reaction zone is operated is operated such that saidreaction zone has an inverse temperature profile. In this respect it isto be appreciated that a reaction zone having an inverse temperatureprofile is a reaction zone in which the inlet reaction temperature tothe reaction zone is lower that the process gas outlet reactiontemperature, namely the inverse of the temperature profile naturallyachieved for an endothermic reaction, such as methane aromatization.This inverse temperature profile can be achieved with countercurrentflow of the feedstock and the particulate dehydrocyclization catalystby, for example, introducing hot catalyst at the top of the reactionzone so that the catalyst moves downward through the reaction zone, withthe reduced temperature catalyst being removed from the bottom of thereaction zone. Feed is introduced at bottom of the reaction zone andflows countercurrent to the catalyst up the reaction zone so that itcontacts the hottest portion of the catalyst at the process gas outlet.

During the methane aromatization reaction, coke tends to build up on theparticulate dehydrocyclization catalyst and hence a portion of theparticulate dehydrocyclization catalyst may be periodically regeneratedin a regeneration zone, which is separate from the reaction zone and isnormally operated under oxidizing conditions. Under the oxidizingconditions in the regeneration zone, coke is burnt from the catalyst butat the same time the activity of the catalyst tends to be adverselyaffected, either by conversion of elemental metal or metal carbides onthe catalyst to oxide forms or by generation of coke selective sites onthe catalyst. Accordingly, in one embodiment of the present process, theregenerated catalyst is transferred to a catalyst treatment zoneseparate from the reaction zone and the regeneration zone, where theregenerated catalyst is contacted with a carburizing gas at atemperature less than the temperature in the reaction zone, butgenerally greater than the temperature in the regeneration zone. The useof the separate catalyst treatment zone allows the contact with thecarburizing gas to be conducted under conditions which favor conversionof metal oxides on the regenerated catalyst back to carbide species orthe elemental form as well as enhancing the aromatics selectivity of thecatalyst. Moreover, any by-products, such as hydrogen, generated as aresult of the contact with the carburizing gas can be removed from thecatalyst treatment zone without being combined with the effluent fromthe reaction zone.

The dehydrocyclization reaction is endothermic and the presentdisclosure also provides a method for supplying heat to the reaction bywithdrawing a further portion of the particulate dehydrocyclizationcatalyst from the reaction zone, heating the further catalyst portion ina heating zone with hot combustion gases generated by burning asupplemental source of fuel and then returning the heated catalystportion to the reaction zone. The heated catalyst portion is preferablyfed to the catalyst treatment zone for contact with the carburizing gasbefore being returned to the reaction zone.

In addition, this disclosure provides a process for utilizing thehydrogen generated as a by-product of the dehydrocyclization reactionand in particular to a process for converting at least part of thehydrogen to higher value products.

Feedstock

Any methane-containing feedstock can be used in the process describedherein but in general the present process is intended for use with anatural gas feedstock. Other suitable methane-containing feedstocksinclude those obtained from sources such as coal beds, landfills,agricultural or municipal waste fermentation, and/or refinery gasstreams.

Methane-containing feedstocks, such as natural gas, typically containcarbon dioxide and ethane in addition to methane. Ethane and otheraliphatic hydrocarbon(s) that may be present in the feed can of coursebe converted to desired aromatics products in the dehydrocyclizationstep. In addition, as will be discussed below, carbon dioxide can alsobe converted to useful aromatics products either directly in thedehydrocyclization step or indirectly through conversion to methaneand/or ethane in the hydrogen rejection step.

Nitrogen and/or sulfur impurities are also typically present inmethane-containing streams may be removed, or reduced to low levels,prior to use of the streams in the process of this disclosure. In anembodiment, the feed to the dehydrocyclization step contains less than100 parts per million by weight (wtppm), for example less than 10 wtppm,such as less than 1 wtppm each of nitrogen and sulfur compounds.

In addition to methane, the feed to the dehydrocyclization step maycontain at least one of hydrogen, water, carbon monoxide and carbondioxide in order to assist in coke mitigation. These additives can beintroduced as separate co-feeds or can be present in the methane stream,such as, for example, where the methane stream is derived from naturalgas containing carbon dioxide. Other sources of carbon dioxide mayinclude flue gases, LNG plants, hydrogen plants, ammonia plants, glycolplants and phthalic anhydride plants.

In some embodiments, the feed to the dehydrocyclization step containscarbon dioxide and comprises about 90 to about 99.9 mol. %, such asabout 97 to about 99 mol. %, methane and about 0.1 to about 10 mol. %,such as about 1 to about 3 mol. %, CO₂. In some additional embodiments,the feed to the dehydrocyclization step contains carbon monoxide andcomprises about 80 to about 99.9 mol. %, such as about 94 to about 99mol. %, methane and about 0.1 to about 20 mol. %, such as about 1 toabout 6 mol. %, CO. In some further embodiments, the feed to thedehydrocyclization step contains steam and comprises about 90 to about99.9 mol. %, such as about 97 to about 99 mol. %, methane and about 0.1to about 10 mol. %, such as about 1 to about 5 mol. %, steam. In yet afurther embodiment, the feed to the dehydrocyclization step containshydrogen and comprises about 80 to about 99.9 mol. %, such as about 95to about 99 mol. %, methane and about 0.1 to about 20 mol. %, such asabout 1 to about 5 mol. %, hydrogen.

The feed to the dehydrocyclization step can also contain higherhydrocarbon(s) than methane, including aromatic hydrocarbon(s). Suchhigher hydrocarbon(s) can be recycled from the hydrogen rejection step,added as separate co-feeds or can be present in the methane stream, suchas, for example, when ethane is present in a natural gas feed. Higherhydrocarbon(s) recycled from the hydrogen rejection step typicallyinclude one-ring aromatics and/or paraffins and olefins havingpredominately 6 or less, such as 5 or less, for example 4 or less,typically 3 or less carbon atoms. In general, the feed to thedehydrocyclization step contains less than 5 wt. %, such as less than 3wt. %, of C₃+ hydrocarbon(s).

Dehydrocyclization

In the dehydrocyclization step of the present process, the methanecontaining feedstock is contacted with a dehydrocyclization catalystunder conditions, normally non-oxidizing conditions and preferablyreducing conditions, effective to convert the methane to higherhydrocarbon(s), including benzene and naphthalene. The principal netreactions involved are as follows:2CH₄

C₂H₄+2H₂  (Reaction 1)6CH₄

C₆H₆+9H₂  (Reaction 2)10CH₄

C₁₀H₈+16H₂  (Reaction 3)

Carbon monoxide and/or dioxide that may be present in the feed improvescatalyst activity and stability by facilitating reactions such as:CO₂+coke

2CO  (Reaction 4)but negatively impacts equilibrium by allowing competing net reactions,such as;CO₂+CH₄

CO+2H₂  (Reaction 5).

Suitable conditions for the dehydrocyclization step include atemperature of about 400° C. to about 1200° C., such as about 500° C. toabout 975° C., for example about 600° C. to about 950° C., a pressure ofabout 1 kPa-a to about 1000 kPa-a, such as about 10 to about 500 kPa-a,for example about 50 kPa-a to about 200 kPa-a and a weight hourly spacevelocity of about 0.01 to about 1000 hr⁻¹, such as about 0.1 to about500 hr⁻¹, for example about 1 to about 20 hr⁻¹. Conveniently, thedehydrocyclization step is conducted in non-oxidizing conditions.Generally, the dehydrocyclization reaction conditions are sufficient toconvert at least 5 wt. %, for example 7 wt. %, such as at least 10 wt.%, for example at least 12 wt. %, and such as at least 15 wt. %, of themethane in the feedstock to higher hydrocarbon(s), generally to aromatichydrocarbon(s), and particularly to benzene.

Any particulate dehydrocyclization catalyst effective to convert methaneto aromatics can be used in the present process, although generally thecatalyst will include a metal component, particularly a transition metalor compound thereof, on an inorganic support. Conveniently, the metalcomponent is present in an amount between about 0.1% and about 20%, suchas between about 1% and about 10%, by weight of the total catalyst.Generally, the metal will be present in the catalyst in elemental formor as a carbide species.

Suitable metal components for the catalyst include calcium, magnesium,barium, yttrium, lanthanum, scandium, cerium, titanium, zirconium,hafnium, vanadium, niobium, tantalum, chromium, molybdenum, tungsten,manganese, rhenium, iron, ruthenium, cobalt, rhodium, iridium, nickel,palladium, copper, silver, gold, zinc, aluminum, gallium, silicon,germanium, indium, tin, lead, bismuth and transuranium metals. Suchmetal components may be present in elemental form or as metal compounds,such as oxides, carbides, nitrides and/or phosphides, and may beemployed alone or in combination. Platinum and osmium can also be usedas one of the metal component but, in general, are not preferred.

The inorganic support may be either amorphous or crystalline and inparticular may be an oxide, carbide or nitride of boron, aluminum,silicon, phosphorous, titanium, scandium, chromium, vanadium, magnesium,manganese, iron, zinc, gallium, germanium, yttrium, zirconium, niobium,molybdenum, indium, tin, barium, lanthanum, hafnium, cerium, tantalum,tungsten, or other transuranium elements. In addition, the support maybe a porous material, such as a microporous crystalline material or amesoporous material. As used herein the term “microporous” refers topores having a diameter of less than 2 nanometers, whereas the term“mesoporous” refers to pores having a diameter of from 2 to 50nanometers.

Suitable microporous crystalline materials include silicates,aluminosilicates, titanosilicates, aluminophosphates, metallophosphates,silicoaluminophosphates or their mixtures. Such microporous crystallinematerials include materials having the framework types MFI (e.g., ZSM-5and silicalite), MEL (e.g., ZSM-11), MTW (e.g., ZSM-12), TON (e.g.,ZSM-22), MTT (e.g., ZSM-23), FER (e.g., ZSM-35), MFS (e.g., ZSM-57), MWW(e.g., MCM-22, PSH-3, SSZ-25, ERB-1, ITQ-1, ITQ-2, MCM-36, MCM-49 andMCM-56), IWR (e.g., ITQ-24), KFI (e.g., ZK-5), BEA (e.g., zeolite beta),ITH (e.g., ITQ-13), MOR (e.g., mordenite), FAU (e.g., zeolites X, Y,ultrastabilized Y and dealuminized Y), LTL (e.g., zeolite L), IWW (e.g.,ITQ-22), VFI (e.g., VPI-5), AEL (e.g., SAPO-11), AFI (e.g., ALPO-5) andAFO (SAPO-41), as well as materials such as MCM-68, EMM-1, EMM-2,ITQ-23, ITQ-24, ITQ-25, ITQ-26, ETS-2, ETS-10, SAPO-17, SAPO-34 andSAPO-35. Suitable mesoporous materials include MCM-41, MCM-48, MCM-50,FSM-16 and SBA-15.

Examples of suitable catalysts include molybdenum, tungsten, zinc,rhenium and compounds and combinations thereof on ZSM-5, silica oralumina.

The metal component can be dispersed on the inorganic support by anymeans well known in the art such as co-precipitation, incipient wetness,evaporation, impregnation, spray-drying, sol-gel, ion-exchange, chemicalvapor deposition, diffusion and physical mixing. In addition, theinorganic support can be modified by known methods, such as, forexample, steaming, acid washing, caustic washing and/or treatment withsilicon-containing compounds, phosphorus-containing compounds, and/orelements or compounds of Groups 1, 2, 3 and 13 of the Periodic Table ofElements (IUPAC 2005). Such modifications can be used to alter thesurface activity of the support and hinder or enhance access to anyinternal pore structure of the support.

In some embodiments, the catalytic particulate material may furthercomprise non-catalytic particulate material. The non-catalyticparticulate material may be used as a material to transport energy(heat) into the system and/or to fill space as required providing therequired hydrodynamic environment. The non-catalytic particulatematerial may form particulates without binder or be bound by aninorganic binder such as clay, silica, alumina, zirconia, or other metaloxide may be used to help maintain the physical integrity of theparticles. Preferably the particles are of a substantially sphericalshape. Examples of suitable non-catalytic particulate material are lowsurface area silica, alumina, ceramics, and silicon carbide.

The dehydrocyclization step is conducted by contacting themethane-containing feedstock with the particulate dehydrocyclizationcatalyst in one or more moving bed reaction zones. Generally, thefeedstock is contacted in each reaction zone with the dehydrocyclizationcatalyst, wherein the bulk of the feedstock flows countercurrent to thedirection of movement of the bulk of the dehydrocyclization catalyst. Insome embodiments, the reaction zone comprises a plurality ofseries-connected moving bed reaction zones in which particulate catalystis cascaded in one direction from one reaction zone to the next adjacentreaction zone in the series, while the feed is passed through andbetween the reaction zones in the opposite direction. In one embodiment,the moving bed reaction zones are fluidizing bed reaction zones.

In one embodiment, the countercurrent flow of the feedstock and theparticulate dehydrocyclization catalyst is arranged to produce aninverse temperature profile in the or each reaction zone, such that,despite the endothermic nature of the dehydrocyclization reaction, thedifference between the process gas outlet reaction temperature from thereaction zone and the inlet reaction temperature to the reaction zone isat least +10° C., such as at least +50° C., for example at least +100°C., and even at least +150° C.

In some additional embodiments, the reaction zone comprises a settlingbed reaction zone. For example, the reaction zone comprises a verticallydisposed reaction zone in which catalytic particulate material enters ator near the top of the reaction zone and flows under gravity to form acatalyst bed, while the feed enters the reaction zone at or near thebase of the reaction zone and flows upwardly through the catalyst bed.

The movement of the dehydrocyclization catalyst in the reaction zone issubstantially free of fluidization in the settling bed embodiment. Theterm “substantially free of fluidization” as used herein means that theaverage gas flowing velocity in the reactor is lower than the minimumfluidizing velocity. The term “substantially free of fluidization” asused herein also means that the average gas flowing velocity in thereactor is less than 99%, such as less than 95%, typically less than90%, even less than 80% of the minimum fluidization velocity. Thecatalytic particulate material contacts the hydrocarbon in asubstantially countercurrent fashion. The term “substantiallycountercurrent fashion” as used herein means the majority of thecatalytic particulate material moves countercurrently to the majority ofthe hydrocarbon.

Where the or each reaction zone is operated as a settling bed, theparticulate catalytic material and/or the particulate non-catalyticmaterial has an average particle size from about 0.1 mm to about 100 mm,such as from about 1 mm to about 5 mm, and for example from about 2 mmto about 4 mm. In some embodiments, at least 90 wt. % of the particulatecatalytic material and/or at least 90 wt. % of the particulatenon-catalytic material has an particle size from about 0.1 mm to about100 mm, such as from about 1 mm to about 5 mm, for example from about 2mm to about 4 mm.

In some embodiments, wherein the reaction zones are operated asfluidizing beds, the catalytic particulate material and/or thenon-catalytic particulate material has an average particle size fromabout 0.01 mm to about 10 mm, such as from about 0.05 mm to about 1 mm,and for example from about 0.1 mm to about 0.6 mm. In some embodiments,at least 90 wt. % of the catalytic particulate material and/or at least90 wt. % of the non-catalytic particulate material have particle sizefrom about 0.01 mm to about 10 mm, such as from about 0.05 to about 1mm, and for example from about 0.1 to about 0.6 mm.

In some embodiments, the mass ratio of the flowrate of the catalyticparticulate material plus any non-catalytic particulate material overthe flowrate of the hydrocarbon feedstock is from about 1:1 to about100:1, such as from about 1:1 to about 40:1, for example from about 5:1to 20:1.

In some embodiments, the reaction zone is located in a cold wallreactor. Operation of the metallic shell at lower temperature than theprocess temperature reduces the cost of the vessel by reducing therequired thickness of the vessel wall as well as potentially enablinguse of lower cost metal alloy.

In some embodiments, the catalytic particulate material enters thereaction zone at a temperature of about 800° C. to about 1200° C., suchas about 900° C. to about 1100° C., and exits the reaction zone at atemperature of about 500° C. to about 800° C., such as about 600° C. toabout 700° C. The total temperature difference of the catalyticparticulate material across the reaction zones is at least 100° C.

The major components of the effluent from the dehydrocyclization stepare hydrogen, benzene, naphthalene, carbon monoxide, ethylene, andunreacted methane. Typically, the effluent contains at least 5 wt. %,such as at least 10 wt. %, for example at least 20 wt. %, convenientlyat least 30 wt. %, more aromatic rings than the feed.

The benzene and naphthalene are then recovered from thedehydrocyclization effluent, for example, by solvent extraction followedby fractionation. However, as will be discussed below, at least part ofthese aromatic components can be submitted to an alkylation step, beforeor after product recovery, to produce higher value materials, such asxylenes.

Catalyst Reheating

The dehydrocyclization reaction is endothermic and in order to supplyheat to the reaction, a first portion of the catalyst may be withdrawnfrom the reaction zone, either on an intermittent, or more preferably, acontinuous basis, and transferred to a separate heating zone, where thefirst catalyst portion is heated by direct contact with hot combustiongases generated by burning a supplemental source of fuel. The heatedfirst catalyst portion is then returned to the reaction zones.

Typically, the supplemental source of fuel comprises a hydrocarbon, suchas methane, and in particular a suitable fuel source is the natural gasused as the feedstock to the process. Conveniently, an oxygen-leanatmosphere is maintained in the heating zone so that burning thehydrocarbon fuel to heat the first catalyst portion produces synthesisgas, which can then be used to generate additional hydrocarbon productand/or fuel. In addition, the use of an oxygen-lean atmosphere inhibitsoxidation of metal carbides present in the dehydrocyclization catalystand minimizes the average steam partial pressure thereby reducingcatalyst hydrothermal aging.

Alternatively, a suitable supplemental fuel source is hydrogen and, inparticular, part of the hydrogen generated as a by-product of thearomatization reaction.

Conveniently, said first catalyst portion is contacted directly with theburning source of fuel in the heating zone. Alternatively, the source offuel is burned in a combustion zone separate from said heating zone andthe combustion gases generated in the combustion zone are fed to theheating zone to heat the first catalyst portion.

In one practical embodiment, the heating zone is elongated and the firstcatalyst portion is passed through the heating zone from an inlet at oradjacent one end of the heating zone to an outlet at or adjacent theother end of the heating zone, with heat being applied to first catalystportion at a plurality of locations spaced along the length of theheating zone. In this way, the heat input to the first catalyst portioncan be distributed along the length of the heating zone therebyminimizing catalyst surface temperatures and internal gradients.

Where the first catalyst portion is heated by direct contact with theburning source of fuel in the heating zone, gradual heating of thecatalyst can be achieved by supplying substantially all of thesupplemental fuel to the inlet end of the heating zone and thensupplying the oxygen-containing gas incrementally to said heating zoneat said plurality of spaced locations along the length of heating zone.Alternatively, substantially all of the oxygen-containing gas requiredto burn said supplemental fuel can be supplied to the inlet end of theheating zone and the supplemental fuel supplied incrementally to theheating zone at said plurality of spaced locations.

Where the first catalyst portion is heated by direct contact with hotcombustion gases generated in a separate combustion zone, gradualheating of the catalyst can be achieved by supplying the hot combustiongases to said plurality of spaced locations along the length of heatingzone.

In some embodiments, the heating zone is a riser and said first catalystportion is passed upwardly through the riser during the reheating step.In practice, the heating zone may include a plurality of risersconnected in parallel. Alternatively, said heating zone can include amoving bed of said catalyst.

Typically, the first catalyst portion is at a temperature of about 500°C. to about 900° C. on entering the heating zone and is at a temperatureof about 800° C. to about 1000° C. on leaving the heating zone. The hotcombustion gases are typically at a temperature of less than 1300° C.,preferably less than 1100° C., more preferably less than 1000° C., forexample at a temperature in the range of about 800° C. to less than1000° C. Typically, the heating zone will be operated at pressuresbetween 10 and 100 psia (69 and 690 kPa-a), more preferably between 15and 60 psia (103 and 414 kPa-a). Typically, the average residence timeof catalyst particles in the heating zone will be between 0.1 and 100seconds, more preferably between 1 and 10 seconds.

Prior to being reintroduced into the reaction zone(s) and, preferablyafter passage through the heating zone, the first catalyst portion maybe subjected to one or more stripping steps to at least partially remove(a) coke or heavy hydrocarbon(s) that may have been produced on thesurface of the catalyst and/or (b) water or oxygen that may have beenadsorbed by the catalyst. Stripping to remove coke or heavyhydrocarbon(s) is conveniently effected by contacting the first catalystportion with steam, hydrogen and/or CO₂, whereas stripping to removewater or oxygen is conveniently effected by contacting the firstcatalyst portion with methane, CO₂ or hydrogen.

In addition, since the reheating step may tend to oxidize catalyticallyactive metal species, particularly metal carbides, contained by thefirst catalyst portion, the reheated catalyst is preferably subjected toa carburizing step prior to being reintroduced into the reaction zone.Conveniently, the carburization step is effected by contacting the firstcatalyst portion with H₂, and CO, CO₂, and/or a hydrocarbon, such asmethane, ethane, or propane, and can be conducted simultaneously with orseparately from the water/oxygen stripping step. Preferably,carburization of the reheated catalyst is effected in the catalysttreatment zone discussed in detail below.

Catalyst Regeneration

As well as being endothermic, the dehydrocyclization reaction tends todeposit coke on the catalyst and hence, to maintain the activity of thedehydrocyclization catalyst, a second portion of the catalyst iswithdrawn from the reaction zone, either on an intermittent, or acontinuous basis, and transferred to a separate regeneration zone. Thegas used to transport the second catalyst portion to the regenerationzone may contain O₂ but preferably contains less O₂ than air, such asless than 10 wt. % O₂, most preferably less than 5% O₂. The transportinggas may contain CO₂ and/or H₂ to gasify a portion of the coke from thesecond catalyst portion, but preferably is substantially free of H₂O andis at a low temperature (typically less than 200° C.) so that thecatalyst stream does not oxidize and heat up above the targettemperature of the regeneration zone.

In the regeneration zone, the second catalyst portion is generallycontacted with an oxygen-containing gas under conditions to at leastpartially remove the coke on the catalyst and thereby regenerate thecatalyst. The regeneration gas preferably contains less O₂ than air,such as less than 10 wt. %, more preferably less than 5 wt. %, O₂, andis preferably substantially free of H₂O. The regeneration gas may alsocontain CO₂ to gasify a portion of the coke from the second catalystportion. Convenient sources of the regeneration gas are an O₂ depleted,N₂ enriched stream from an air separation unit and a high CO₂ rejectstream from industrial or natural gas processing to which air or O₂ hasbeen added to achieve the target O₂ concentration. Typically theregeneration gas is circulated between the regeneration zone and aconditioning zone, where the used regeneration gas is cooled to condenseout excess water, make-up oxygen-containing gas (preferably air) isadded to maintain the target O₂ concentration and a portion is purged tomaintain constant pressure. Typically the regeneration zone will beoperated at pressures between 10 and 100 psia (69 and 690 kPa-a), morepreferably between 15 and 60 psia (103 and 414 kPa-a).

The regeneration zone may be a reactor operated as a fluidizing bed, anebulating bed, a settling bed, a riser reactor or a combination thereof.In practice, the regeneration zone may include a plurality of reactors,such as a plurality of riser reactors connected in parallel. Theregeneration zone should be operated at the minimum temperature requiredto remove the required amount of coke at the design residence time andin particular the temperature should not exceed the point at which metaloxide volatilization occurs or the catalyst substrate undergoes rapiddeterioration. Generally, the temperature in the regeneration zone isless than the temperature of the reaction zone and typicallyregeneration zone temperature is from about 400° C. to about 700° C.,such as from about 550° C. to about 650° C. Catalyst residence time inthe regeneration zone also should be minimized to reduce catalyst agingrate and maximize percent of time the catalyst spends in the reactordoing useful work. Typically, the average residence time of catalystparticles in the regeneration zone will be between 0.1 and 100 minutes,more preferably between 1 and 20 minutes.

Conveniently, the ratio of the weight of the first catalyst portiontransferred in a given time to the heating zone to the weight of secondcatalyst portion transferred in the same time to the regeneration zoneis in the range of about 5:1 to about 100:1, preferably about 10:1 toabout 20:1.

In addition to removing coke on the catalyst, the oxygen-containing gasin the regeneration zone tends to react with the metal on the catalyst,thereby converting the metal from the elemental or carbidic speciesdesired for the dehydroaromatization reaction to less active oxidespecies. Moreover, and particularly where the support is a zeolite, theregeneration step may produce active sites on the surface of thecatalyst support that favor coke deposition. Thus, before being returnedto the reaction zone, the regenerated catalyst is transferred to acatalyst treatment zone separate from the regeneration zone, the heatingzone and the reaction zone, where the regenerated catalyst is contactedwith a carburizing gas containing at least one hydrocarbon selected frommethane, ethane, propane, butane, isobutene, benzene and naphthalene. Insome cases, the carburizing gas may also contain at least one of CO₂,CO, H₂, H₂O and other diluents. Moreover, it may be desirable to contactthe regenerated catalyst sequentially with a plurality of differenthydrocarbon(s), each hydrocarbon being selected from methane, ethane,propane, butane, isobutene, benzene and naphthalene.

The catalyst treatment zone may be operated as a fluidizing bed reactor,ebulating bed reactor, settling bed reactor, riser reactor orcirculating riser reactor. In one preferred embodiment, the catalysttreatment zone comprises a settling bed reactor. Alternatively, thecatalyst treatment zone comprises a single fluidizing bed reactor withinternal baffles to prevent back-mixing or a plurality of fluidizing bedreactors in series with the regenerated catalyst being cascaded betweenadjacent reactors. In any event, contact in the catalyst treatment zoneis facilitated by arranging that the regenerated catalyst and thecarburizing gas flow in opposite directions in said catalyst treatmentzone.

For some catalysts, it may be preferable that the regenerated catalystportion is initially contacted with a H₂-rich stream to partially orfully reduce the metal component of the catalyst prior to thecarburization step. It may also be desirable to subject the carburizedcatalyst to post treatment with H₂ and/or CO₂ to strip off any excesscarbon that may have been deposited on the catalyst by the carburizationstep.

After leaving the carburization zone, the second catalyst portion isreturned to the reaction zone to contact the methane feed. In onepractical embodiment, the dehydrocyclization step is conducted invertically-disposed, settling bed reactors with the feedstock enteringthe lower reactor at or near the base of the reactor and the heatedfirst catalyst portion and the regenerated second catalyst portion beingreturned to the upper reactor at or near the top of the reactor.Conveniently, the hydrocarbon effluent is recovered from at or near thetop of the lower reactor and conveyed to enter the upper reactor at ornear the base. Conveniently, said first and second catalyst portions areremoved from at or near the base of the upper reactor and conveyed toenter at or near the top of the lower reactor and the process effluentis recovered from at or near the top of the upper reactor.

In an alternative embodiment, the dehydrocyclization step is conductedin a plurality of fluidizing bed reactors connected in series, with thefeedstock entering the first reactor in the series and the heated firstcatalyst portion and the regenerated second catalyst portion beingreturned to the final reactor in the series. The hydrocarbon stream andcatalyst particulate stream are then conveyed counter current to oneanother through the series of reactors. Conveniently, said first andsecond catalyst portions are removed from the first reactor.

In yet a further embodiment, the regeneration or catalyst coke strippingmay be effected utilizing hydrogen containing gas. The regenerationconditions when utilizing hydrogen comprise a temperature from about600° C. to about 1000° C.; such as from about 700° C. to about 950° C.,for example from about 800° C. to about 900° C. Generally the hydrogencontaining gas should not contain significant quantities of methane orother hydrocarbons; and typically contains less than 20 mol %; such asless than 10 mol %; for example less than 2 mol % hydrocarbon.

Hydrogen Management

Since hydrogen is a major component of the dehydrocyclization effluent,after recovery of the aromatic products, the effluent is subjected to ahydrogen rejection step to reduce the hydrogen content of the effluentbefore the unreacted methane is recycled to the dehydrocyclization stepand to maximize feed utilization. Typically the hydrogen rejection stepcomprises reacting at least part of the hydrogen in thedehydrocyclization effluent with an oxygen-containing species,preferably CO and/or CO₂, to produce water and a second effluent streamhaving a reduced hydrogen content compared with the first(dehydrocyclization) effluent stream.

Conveniently, the hydrogen rejection step includes (i) methanationand/or ethanation, (ii) a Fischer-Tropsch process, (iii) synthesis of C₁to C₃ alcohols, particularly methanol, and other oxygenates, (iv)synthesis of light olefins, paraffins and/or aromatics by way of amethanol or dimethyl ether intermediate and/or (v) selective hydrogencombustion. These steps may be employed sequentially to gain thegreatest benefit; for example Fischer-Tropsch may first be employed toyield a C₂+ enriched stream followed by methanation to achieve highconversion of the H₂.

Typically, as described below, the hydrogen rejection step will generatehydrocarbon(s), in which case, after separation of the co-producedwater, at least portion of the hydrocarbon(s) are conveniently recycledto the dehydrocyclization step. For example, where the hydrocarbon(s)produced in the hydrogen rejection step comprise paraffins and olefins,the portion recycled to the dehydrocyclization step convenientlycomprises, paraffins or olefins with 6 or less carbon atoms, such as 5or less carbon atoms, for example 4 or less carbon atoms or 3 or lesscarbon atoms. Where, the hydrocarbon(s) produced in the hydrogenrejection step comprise aromatics, the portion recycled to thedehydrocyclization step conveniently comprises single ring aromaticspecies.

Methanation/Ethanation

In some embodiments the hydrogen rejection step comprises reaction of atleast part of the hydrogen in the dehydrocyclization effluent withcarbon dioxide to produce methane and/or ethane according to thefollowing net reactions:CO₂+4H₂

CH₄+2H₂O  (Reaction 6)2CO₂+7H₂

C₂H₆+4H₂O  (Reaction 7)

The carbon dioxide employed is conveniently part of a natural gas streamand preferably the same natural gas stream used as the feed to thedehydrocyclization step. Where the carbon dioxide is part of amethane-containing stream, the CO₂:CH₄ of the stream is convenientlymaintained between about 1:1 and about 0.1:1. Mixing of the carbondioxide-containing stream and the dehydrocyclization effluent isconveniently achieved by supplying the gaseous feeds to the inlet of ajet ejector.

The hydrogen rejection step to produce methane or ethane normallyemploys a H₂:CO₂ molar ratio close to the stoichiometric proportionsrequired for the desired Reaction 6 or Reaction 7, although smallvariations can be made in the stoichiometric ratio if it is desired toproduce a CO₂-containing or H₂-containing second effluent stream. Thehydrogen rejection step to produce methane or ethane is convenientlyeffected in the presence of a bifunctional catalyst comprising a metalcomponent, particularly a transition metal or compound thereof, on aninorganic support. Suitable metal components comprise copper, iron,vanadium, chromium, zinc, gallium, nickel, cobalt, molybdenum,ruthenium, rhodium, palladium, silver, rhenium, tungsten, iridium,platinum, gold, gallium and combinations and compounds thereof. Theinorganic support may be an amorphous material, such as silica, aluminaor silica-alumina, or like those listed for the dehydroaromatizationcatalyst. In addition, the inorganic support may be a crystallinematerial, such as a microporous or mesoporous crystalline material.Suitable porous crystalline materials include the aluminosilicates,aluminophosphates and silicoaluminophosphates listed above for thedehydrocyclization catalyst.

The hydrogen rejection step to produce methane and/or ethane can beconducted over a wide range of conditions including a temperature ofabout 100° C. to about 900° C., such as about 150° C. to about 500° C.,for example about 200° C. to about 400° C., a pressure of about 200kPa-a to about 20,000 kPa-a, such as about 500 to about 5000 kPa-a and aweight hourly space velocity of about 0.1 to about 10,000 hr⁻¹, such asabout 1 to about 1,000 hr⁻¹. CO₂ conversion levels are typically between20 and 100% and preferably greater than 90%, such as greater than 99%.This exothermic reaction may be carried out in multiple catalyst bedswith heat removal between beds. In addition, the lead bed(s) may beoperated at higher temperatures to maximize kinetic rates and the tailbeds(s) may be operated at lower temperatures to maximize thermodynamicconversion.

The main products of the reaction are water and, depending on the H₂:CO₂molar ratio, methane, ethane and higher alkanes, together with someunsaturated C₂ and higher hydrocarbon(s). In addition, some partialhydrogenation of the carbon dioxide to carbon monoxide is preferred.After removal of the water, the methane, carbon monoxide, any unreactedcarbon dioxide and higher hydrocarbon(s) can be fed directly to thedehydrocyclization step to generate additional aromatic products.

Fischer-Tropsch Process

In some additional embodiments the hydrogen rejection step comprisesreaction of at least part of the hydrogen in the dehydrocyclizationeffluent with carbon monoxide according to the Fischer-Tropsch processto produce C₂ to C₅ paraffins and olefins.

The Fischer-Tropsch process is well known in the art, see for example,U.S. Pat. Nos. 5,348,982 and 5,545,674 incorporated herein by reference.The process typically involves the reaction of hydrogen and carbonmonoxide in a molar ratio of about 0.5:1 to about 4:1, preferably about1.5:1 to about 2.5:1, at a temperature of about 175° C. to about 400°C., preferably about 180° C. to about 240° C. and a pressure of about 1to about 100 bar (100 to 10,000 kPa-a), preferably about 10 to about 40bar (1,000 to 4,000 kPa-a), in the presence of a Fischer-Tropschcatalyst, generally a supported or unsupported Group VIII, non-noblemetal, e.g., Fe, Ni, Ru, Co, with or without a promoter, e.g. ruthenium,rhenium, hafnium, zirconium, titanium. Supports, when used, can berefractory metal oxides such as Group IVB, i.e., titania, zirconia, orsilica, alumina, or silica-alumina. In some embodiments, the catalystcomprises a non-shifting catalyst, e.g., cobalt or ruthenium, preferablycobalt, with rhenium or zirconium as a promoter, preferably cobalt andrhenium supported on silica or titania, preferably titania.

In some additional embodiments, the hydrocarbon synthesis catalystcomprises a metal, such as Cu, Cu/Zn or Cr/Zn, on the ZSM-5 and theprocess is operated to generate significant quantities of single-ringaromatic hydrocarbon(s). An example of such a process is described inStudy of Physical Mixtures of Cr ₂0₃-ZnO and ZSM-5 Catalysts for theTransformation of Syngas into Liquid Hydrocarbon(s) by Jose Erena; Ind.Eng. Chem. Res. 1998, 37, 1211-1219, incorporated herein by reference.

The Fischer-Tropsch liquids, i.e., C₅+, are recovered and light gases,e.g., unreacted hydrogen and CO, C₁ to C₃ or C₄ and water are separatedfrom the heavier hydrocarbon(s). The heavier hydrocarbon(s) can then berecovered as products or fed to the dehydrocyclization step to generateadditional aromatic products.

The carbon monoxide required for the Fischer-Tropsch reaction can beprovided wholly or partly by the carbon monoxide present in or cofedwith the methane-containing feed and generated as a by-product in thedehydrocyclization step. If required, additional carbon monoxide can begenerated by feeding carbon dioxide contained, for example, in naturalgas, to a shift catalyst whereby carbon monoxide is produced by thereverse water gas shift reaction:CO₂+H₂

CO+H₂O  (Reaction 8)and by the following reaction:CH₄+H₂O

CO+3H₂Alcohol Synthesis

In some further embodiments the hydrogen rejection step comprisesreaction of at least part of the hydrogen in the dehydrocyclizationeffluent with carbon monoxide to produce C₁ to C₃ alcohols, andparticularly methanol. The production of methanol and other oxygenatesfrom synthesis gas is also well-known and is described in, for example,in U.S. Pat. Nos. 6,114,279; 6,054,497; 5,767,039; 5,045,520; 5,254,520;5,610,202; 4,666,945; 4,455,394; 4,565,803; 5,385,949, the descriptionsof which are incorporated herein by reference. Typically, the synthesisgas employed has a molar ratio of hydrogen (H₂) to carbon oxides(CO+CO₂) in the range of from about 0.5:1 to about 20:1, preferably inthe range of from about 2:1 to about 10:1, with carbon dioxideoptionally being present in an amount of not greater than 50% by weight,based on total weight of the syngas.

The catalyst used in the methanol synthesis process generally includesan oxide of at least one element selected from the group consisting ofcopper, silver, zinc, boron, magnesium, aluminum, vanadium, chromium,manganese, gallium, palladium, osmium and zirconium. Conveniently, thecatalyst is a copper based catalyst, such as in the form of copperoxide, optionally in the presence of an oxide of at least one elementselected from silver, zinc, boron, magnesium, aluminum, vanadium,chromium, manganese, gallium, palladium, osmium and zirconium.Conveniently, the catalyst contains copper oxide and an oxide of atleast one element selected from zinc, magnesium, aluminum, chromium, andzirconium. In some embodiments, the methanol synthesis catalyst isselected from the group consisting of: copper oxides, zinc oxides andaluminum oxides. More preferably, the catalyst contains oxides of copperand zinc.

The methanol synthesis process can be conducted over a wide range oftemperatures and pressures. Suitable temperatures are in the range offrom about 150° C. to about 450° C., such as from about 175° C. to about350° C., for example from about 200° C. to about 300° C. Suitablepressures are in the range of from about 1,500 kPa-a to about 12,500kPa-a, such as from about 2,000 kPa-a to about 10,000 kPa-a, for example2,500 kPa-a to about 7,500 kPa-a. Gas hourly space velocities varydepending upon the type of process that is used, but generally the gashourly space velocity of flow of gas through the catalyst bed is in therange of from about 50 hr⁻¹ to about 50,000 hr⁻¹, such as from about 250hr⁻¹ to about 25,000 hr⁻¹, more preferably from about 500 hr⁻¹ to about10,000 hr⁻¹. This exothermic reaction may be carried out in either fixedor fluidizing beds, including multiple catalyst beds with heat removalbetween beds. In addition, the lead bed(s) may be operated at highertemperatures to maximize kinetic rates and the tail beds(s) may beoperated at lower temperatures to maximize thermodynamic conversion.

The resultant methanol and/or other oxygenates can be sold as a separateproduct, can be used to alkylate the aromatics generated in thedehydrocyclization step to higher value products, such as xylenes, orcan be used as a feedstock for the production of lower olefins,particularly ethylene and propylene. The conversion of methanol toolefins is a well-known process and is, for example, described in U.S.Pat. No. 4,499,327, incorporated herein by reference.

Selective Hydrogen Combustion

In some additional embodiments, the hydrogen rejection step comprisesselective hydrogen combustion, which is a process in which hydrogen in amixed stream is reacted with oxygen to form water or steam withoutsubstantially reacting hydrocarbon(s) in the stream with oxygen to formcarbon monoxide, carbon dioxide, and/or oxygenated hydrocarbon(s).Generally, selective hydrogen combustion is carried out in the presenceof an oxygen-containing solid material, such as a mixed metal oxide,that will release a portion of the bound oxygen to the hydrogen.

One suitable selective hydrogen combustion process is described in U.S.Pat. No. 5,430,210, incorporated herein by reference, and comprisescontacting at reactive conditions a hydrocarbon feedstock comprisinghydrocarbon and hydrogen and a catalytic particulate material comprisingoxygen with separate surfaces of a membrane impervious to non-oxygencontaining gases, wherein said membrane comprises a metal oxideselective for hydrogen combustion, and recovering selective hydrogencombustion product. The metal oxide is typically a mixed metal oxide ofbismuth, indium, antimony, thallium and/or zinc.

U.S. Pat. No. 5,527,979, incorporated herein by reference, describes aprocess for the net catalytic oxidative dehydrogenation of alkanes toproduce alkenes. The process involves simultaneous equilibriumdehydrogenation of alkanes to alkenes and the selective combustion ofthe hydrogen formed to drive the equilibrium dehydrogenation reactionfurther to the product alkenes. In particular, the alkane feed isdehydrogenated over an equilibrium dehydrogenation catalyst in a firstreactor, and the effluent from the first reactor, along with oxygen, isthen passed into a second reactor containing a metal oxide catalystwhich serves to selectively catalyze the combustion of hydrogen. Theequilibrium dehydrogenation catalyst may comprise platinum and theselective metal oxide combustion catalyst may contain bismuth, antimony,indium, zinc, thallium, lead and tellurium or a mixture thereof.

U.S. Patent Application Publication No. 2004/0152586, published Aug. 5,2004 and incorporated herein by reference, describes a process forreducing the hydrogen content of the effluent from a cracking reactor.The process employs a catalyst system comprising (1) at least one solidacid cracking component and (2) at least one metal-based selectivehydrogen combustion component consisting essentially of (a) a metalcombination selected from the group consisting of:

-   -   i) at least one metal from Group 3 and at least one metal from        Groups 4-15 of the Periodic Table of the Elements;    -   ii) at least one metal from Groups 5-15 of the Periodic Table of        the Elements, and at least one metal from at least one of Groups        1, 2, and 4 of the Periodic Table of the Elements;    -   iii) at least one metal from Groups 1-2, at least one metal from        Group 3, and at least one metal from Groups 4-15 of the Periodic        Table of the Elements; and    -   iv) two or more metals from Groups 4-15 of the Periodic Table of        the Elements        and (b) at least one of oxygen and sulfur, wherein the at least        one of oxygen and sulfur is chemically bound both within and        between the metals.

The selective hydrogen combustion reaction of the present disclosure isgenerally conducted at a temperature in the range of from about 300° C.to about 850° C. and a pressure in the range of from about 1 atm toabout 20 atm (100 to 2000 kPa-a).

Aromatic Product Recovery/Treatment

The major products of the dehydrocyclization step are benzene andnaphthalene. These products can be separated from the dehydrocyclizationeffluent, typically by solvent extraction followed by fractionation, andthen sold directly as commodity chemicals. Alternatively, some or all ofthe benzene and/or naphthalene can be alkylated to produce, for example,toluene, xylenes and alkyl naphthalenes and/or can be subjected tohydrogenation to produce, for example, cyclohexane, cyclohexene,dihydronaphthalene (benzylcyclohexene), tetrahydronaphthalene(tetralin), hexahydronaphthalene (dicyclohexene), octahydronaphthaleneand/or decahydronaphthalene (decalin).

Aromatics Alkylation

Alkylation of aromatic compounds such as benzene and naphthalene is wellknown in the art and typically involves reaction of an olefin, alcoholor alkyl halide with the aromatic species in the gas or liquid phase inthe presence of an acid catalyst. Suitable acid catalysts include mediumpore zeolites (i.e., those having a Constraint Index of 2-12 as definedin U.S. Pat. No. 4,016,218), including materials having the frameworktypes MFI (e.g., ZSM-5 and silicalite), MEL (e.g., ZSM-11), MTW (e.g.,ZSM-12), TON (e.g., ZSM-22), MTT (e.g., ZSM-23), MFS (e.g., ZSM-57) andFER (e.g., ZSM-35) and ZSM-48, as well as large pore zeolites (i.e.,those having a Constraint Index of less than 2) such as materials havingthe framework types BEA (e.g., zeolite beta), FAU (e.g., ZSM-3, ZSM-20,zeolites X, Y, ultrastabilized Y and dealuminized Y), MOR (e.g.,mordenite), MAZ (e.g., ZSM-4), MEI (e.g., ZSM-18) and MWW (e.g., MCM-22,PSH-3, SSZ-25, ERB-1, ITQ-1, ITQ-2, MCM-36, MCM-49 and MCM-56).

In some embodiments of the present process, benzene is recovered fromthe dehydrocyclization effluent and then alkylated with an olefin, suchas ethylene produced as a by-product of a hydrogen rejection stepemploying ethanation/methanation. Typical conditions for carrying outthe vapor phase alkylation of benzene with ethylene include atemperature of from about 650 to 900° F. (343 to 482° C.), a pressure ofabout atmospheric to about 3000 psig (100 to 20,800 kPa-a), a WHSV basedon ethylene of from about 0.5 to about 2.0 hr⁻¹ and a mole ratio ofbenzene to ethylene of from 1:1 to 30:1. Liquid phase alkylation ofbenzene with ethylene may be carried out at a temperature between 300and 650° F. (150 to 340° C.), a pressure up to about 3000 psig (20,800kPa-a), a WHSV based on ethylene of from about 0.1 to about 20 hr⁻¹ anda mole ratio of benzene to ethylene of from 1:1 to 30:1.

Preferably, the benzene ethylation is conducted under at least partialliquid phase conditions using a catalyst comprising at least one ofzeolite beta, zeolite Y, MCM-22, PSH-3, SSZ-25, ERB-1, ITQ-1, ITQ-2,ITQ-13, ZSM-5 MCM-36, MCM-49 and MCM-56.

The benzene ethylation can be conducted at the site of thedehydrocyclization/hydrogen rejection process or the benzene can beshipped to another location for conversion to ethylbenzene. Theresultant ethylbenzene can then be sold, used as a precursor in, forexample, the production of styrene or isomerized by methods well knownin the art to mixed xylenes.

In some additional embodiments of the present process, the alkylatingagent is methanol or dimethylether (DME) and is used to alkylate benzeneand/or naphthalene recovered from the dehydrocyclization effluent toproduce toluene, xylenes, methylnaphthalenes and/ordimethylnaphthalenes. Where the methanol or DME is used to alkylatebenzene, this is conveniently effected in presence of catalystcomprising a zeolite, such as ZSM-5, zeolite beta, ITQ-13, MCM-22,MCM-49, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35, and ZSM-48, which hasbeen modified by steaming so as to have a Diffusion Parameter for 2,2dimethylbutane of about 0.1-15 sec⁻¹ when measured at a temperature of120° C. and a 2,2 dimethylbutane pressure of 60 torr (8 kPa-a). Such aprocess is selective to the production of para-xylene and is describedin, for example, U.S. Pat. No. 6,504,272, incorporated herein byreference. Where the methanol is used to alkylate naphthalene, this isconveniently effected in the presence of a catalyst comprising ZSM-5,MCM-22, PSH-3, SSZ-25, ERB-1, ITQ-1, ITQ-2, ITQ-13, MCM-36, MCM-49 orMCM-56. Such a process can be used to selectively produce2,6-dimethylnaphthalene and is described in, for example, U.S. Pat. Nos.4,795,847 and 5,001,295, incorporated herein by reference.

Where methanol or DME is used as an alkylating agent in the process ofthis disclosure, it can be provided as a separate feed to the process orcan at least partly be generated in situ by adding a carbondioxide-containing feed gas, such as a natural gas stream, to part orall of the effluent from the dehydrocyclization step. In particular, thedehydrocyclization effluent, prior to any separation of the aromaticcomponents, can be fed to a reverse shift reactor and reacted with thecarbon dioxide-containing feed under conditions to increase the carbonmonoxide content of the effluent by reactions, such as Reactions 5 and 8above.

In addition, methane and CO₂ and/or steam may be fed to a reverse shiftreactor to generate syngas which can then be mixed with a portion of thedehydrocyclization effluent to adjust the H₂/CO/CO₂ ratios as requiredfor the alkylation step.

Typically, the reverse shift reactor contains a catalyst comprising atransition metal on a support, such as Fe, Ni, Cr, Zn on alumina, silicaor titania, and is operated under conditions including a temperature ofabout 500° C. to about 1200° C., such as about 600° C. to about 1000°C., for example about 700° C. to about 950° C. and a pressure of about 1kPa-a to about 10,000 kPa-a, such as about 2,000 kPa-a to about 10,000kPa-a, for example about 3000 kPa-a to about 5,000 kPa-a. Gas hourlyspace velocities may vary depending upon the type of process used, butgenerally the gas hourly space velocity of flow of gas through thecatalyst bed is in the range of about 50 hr⁻¹ to about 50,000 hr⁻¹, suchas about 250 hr⁻¹ to about 25,000 hr⁻¹, more preferably about 500 hr⁻¹to about 10,000 hr⁻¹.

The effluent from the reverse shift reactor can then be fed to analkylation reactor operating under conditions to cause reactions such asthe following to occur:CO+2H₂

CH₃OH  (Reaction 9)CH₃OH+C₆H₆

toluene+H₂O  (Reaction 10)2CH₃OH+C₆H₆

xylenes+2H₂O  (Reaction 11)

Suitable conditions for such an alkylation reactor would include atemperature of about 100 to about 700° C., a pressure of about 1 toabout 300 atmospheres (100 to 30,000 kPa-a), and a WHSV for the aromatichydrocarbon of about 0.01 to about 100 hr⁻¹. A suitable catalyst wouldcomprise a molecular sieve having a constraint index of 1 to 12, such asZSM-5, typically together with one or metals or metal oxides, such ascopper, chromium and/or zinc oxide.

Preferably, where the alkylation catalyst includes a molecular sieve,the latter is modified to change its diffusion characteristics such thatthe predominant xylene isomer produced by Reaction 11 is paraxylene.Suitable means of diffusion modification include steaming and ex-situ orin-situ deposition of silicon compounds, coke, metal oxides, such asMgO, and/or P on the surface or in the pore mouths of the molecularsieve. Also preferred is that an active metal be incorporated into themolecular sieve so as to saturate more highly reactive species, such asolefins, which may be generated as by-products and which could otherwisecause catalyst deactivation.

The effluent from the alkylation reactor could then be fed to aseparation section in which the aromatic products would initially beseparated from the hydrogen and other low molecular weight materials,conveniently by solvent extraction. The aromatics products could then befractionated into a benzene fraction, a toluene fraction, a C₈ fractionand a heavy fraction containing naphthalene and alkylated naphthalenes.The C₈ aromatic fraction could then be fed to a crystallization orsorption process to separate the valuable p-xylene component and theremaining mixed xylenes either sold as product or fed to anisomerization loop to generate more p-xylene. The toluene fraction couldeither be removed as saleable product, recycled to the alkylationreactor or fed to a toluene disproportionation unit, and preferably aselective toluene disproportionation unit for the preparation ofadditional p-xylene.

Aromatics Hydrogenation

In addition to or instead of the alkylation step, at least part of thearomatic components in the dehydrocyclization effluent can behydrogenated to generate useful products such as cyclohexane,cyclohexene, dihydronaphthalene (benzylcyclohexene),tetrahydronaphthalene (tetralin), hexahydronaphthalene (dicyclohexene),octahydronaphthalene and/or decahydronaphthalene (decalin). Theseproducts can be employed as fuels and chemical intermediates and, in thecase of tetralin and decalin, can be used as the solvent for extractingthe aromatic components from the dehydrocyclization effluent.

The hydrogenation is conveniently, but not necessarily, conducted afterseparation of the aromatic components from the dehydrocyclizationeffluent and conveniently employs part of the hydrogen generated by thedehydrocyclization reaction. Suitable aromatic hydrogenation processesare well known in the art and typically employ a catalyst comprising Ni,Pd, Pt, Ni/Mo or sulfided Ni/Mo supported on alumina or silica support.Suitable operating conditions for the hydrogenation process include atemperature of about 300 to about 1,000° F. (150 to 540° C.), such asabout 500 to about 700° F. (260 to 370° C.), a pressure of about 50 toabout 2,000 psig (445 to 13890 kPa-a), such as about 100 to about 500psig (790 to 3550 kPa-a) and a WHSV of about 0.5 to about 50 hr⁻¹, suchas about 2 to about 10 hr⁻¹.

Partial hydrogenation to leave one or more olefinic carbon-carbon bondsin the product may also be desirable so as to produce materials suitablefor polymerization or other downstream chemical conversion. Suitablepartial hydrogenation processes are well known in the art and typicallyemploy a catalyst comprising noble metals with ruthenium being preferredsupported on metallic oxides, such as La₂O₃—ZnO. Homogeneous noble metalcatalyst systems can also be used. Examples of partial hydrogenationprocesses are disclosed in U.S. Pat. Nos. 4,678,861; 4,734,536;5,457,251; 5,656,761; 5,969,202; and 5,973,218, the entire contents ofwhich are incorporated herein by reference.

An alternative hydrogenation process involves low pressure hydrocrackingof the naphthalene component to produce alkylbenzenes over a catalystsuch as sulfided Ni/W or sulfided Ni supported on an amorphousaluminosilicate or a zeolite, such as zeolite X, zeolite Y or zeolitebeta. Suitable operating conditions for low pressure hydrocrackinginclude a temperature of about 300 to about 1,000° F. (150 to 540° C.),such as about 500 to about 700° F. (260 to 370° C.), a pressure of about50 to about 2,000 psig (445 to 13890 kPa-a), such as about 100 to about500 psig (790 to 3550 kPa-a) and a WHSV of about 0.5 to about 50 hr⁻¹,such as about 2 to about 10 hr⁻¹.

This disclosure will now be more particularly described with referenceto the accompanying drawings and the following non-limiting Examples.

Referring to FIG. 1, the drawing illustrates a simplified design of adehydrocyclization reactor system for converting methane to aromaticsaccording to one embodiment of this disclosure. In this embodiment, thedehydrocyclization reactor includes a settling bed reaction zone, 11, inwhich catalytic particulate material is moved from top of the reactionzone to the bottom of the reaction zone, while the feed is passedthrough the reaction zone in the opposite direction. The heatedcatalytic particulate material flows through an inlet located adjacentthe top of the reactor 11 via line 12. The cooled catalytic particulatematerial flows out of the reactor 11 via outlets located adjacent thebase of the reactor 11 and withdrawn via lines 13 and 14. Methane feedis introduced into the reactor 11 adjacent the base thereof via line 15.The product and unreacted methane flows out of reactor 11 via outlet 16adjacent to the top of reactor 11. Typically, the heated catalyst entersthe reactor 11 at a temperature of about 850° C. and the cooled catalystleaves the reactor at a temperature of about 600° C. FIG. 1 portrays thereactor 11 being one reaction zone. However one having ordinary skill inthe art understands that the reactor system may contain more than morethan one zone.

One having ordinary skill in the art understands that the embodimentsdiscussed in this application do not represent all the possibleapparatus or process variations embodied by the present disclosure. Inaddition, many pieces of equipment and apparatus and certain processingsteps may be needed for industrial, commercial or even experimentalpurposes. Examples of such equipments and apparatus and processing stepsare, but not limited to, distillation columns, fractionation columns,heat exchanges, pumps, valves, pressure gauges, temperature gauges,liquid-vapor separators, feed and product driers and/or treaters, claytreaters, feed and/or product storage facilities, and processes andsteps for process control. While such equipment, apparatus and stepsthat are not needed for understanding the essence of the presentdisclosure are not shown in the drawings, some of them may be mentionedfrom time to time to illustrate various aspects of this disclosure. Itis also noted that some of the equipment may be placed at differentplaces in the process depending on the conditions of the processes.

The invention will now be more particularly described with reference tothe following Examples, which are intended to illustrate the two keybenefits of the process described herein:

(a) Improved product selectivity: reduced production of coke andincreased production of high value products (benzene, toluene, andnaphthalene). Reduction in coke selectivity has two major benefits,namely improved feed utilization and reduction in deactivation rate ofthe catalyst.

(b) Reduced catalyst circulation rates which will reduce the attritionrate of the catalyst as well as reduce the erosion rate of reactionvessels, reactor internals, transfer lines, and other associatedequipment.

EXAMPLE 1

Mo/ZSM-5 catalysts were prepared via impregnation of required amount ofammonium heptamolybdate solution onto NH₄ZSM-5 support (having a Si/Al₂ratio of 28) via incipient wetness, followed by drying at 120° C. for 2hours and final calcination at 500° C. for 6 hours in flowing air. Anominal molybdenum loading (wt. % of metal based on the total weight ofthe catalyst) was targeted of 2.7 wt. %; minor variations in molybdenumloadings do not affect the conclusions obtained. Each Mo/ZSM-5 catalystsample (after calcination) was pelletized, crushed and sieved to 30-60mesh particle size. Catalytic testing of the Mo/ZSM-5 catalysts wasperformed in a quartz reactor packed to form a fixed-bed using quartzwool supports.

Catalyst performance for methane dehydrocyclization to benzene wasperformed at various temperatures using a 95 wt. % CH₄-5 wt. % argonfeed (argon is used as internal standard) at a weight-hourly spacevelocity (based on methane) of 1.2 hr⁻¹. All experimental data wasobtained at 138 kPa-a (20 psia) and all modeling was also performed atthe same pressure. The reaction effluent was analyzed using a massspectrometer and gas chromatograph to determine the methane, benzene,toluene, ethylene, naphthalene, hydrogen, and argon concentrations. Therate of coke deposition on the catalyst (i.e., heavy carbonaceousdeposit which does not volatize from catalyst surface) was determinedvia carbon balance. Additional data was obtained at two temperatures(750° C. and 800° C.) with H₂ added to the feed at 6 mol % and 20 mol %respectively.

For the purposes of these examples the experimental data wasconsolidated to two values Sel_(BTN) and Sel_(Coke). The Sel_(BTN) isthe average selectivity on a carbon molar basis as defined by the sum ofthe moles of carbon in the product present in benzene, toluene, andnaphthalene divided by the moles of carbon contained in methane thatreacted. The Sel_(Coke) is the average selectivity on a carbon molarbasis as defined by the sum of the moles of carbon that remains in thereactor divided by the moles of carbon contained in methane thatreacted. The sum of Sel_(BTN) and Sel_(Coke) does not equal to 100% dueto the formation of other minor products, predominately ethylene. As itis often difficult to obtain accurate experimental thermodynamicconversion data, commercially available simulation software (PROII/6.0Copyright 2003 Invensys Systems Inc.) was utilized to establish thevalue Conv_(BL). The Conv_(BL) is defined as the maximumthermodynamically obtainable conversion of methane to benzene andhydrogen (i.e., no model constrained so that no other products such ascoke, naphthalene, ethylene, etc) at a given temperature and 138 kPa-a(20 psia) pressure. The experimental and modeling results are shown inTable 1.

TABLE 1 Temp H₂ Co- Sel_(BTN) Sel_(Coke) Conv_(BL) ° C. Feed Mol % % Con Feed 600 0% 99% 0.01%   5% 650 0% 98% 0.1%   8% 700 0% 96% 1% 12% 7500% 85% 9% 17% 750 6% 89% 5% 800 0% 68% 24%  23% 800 20%  84% 8% 850 0%45% 46%  29% 900 0% 20% 71%  37%

It is understood that different catalyst compositions, the use ofco-feeds (CO₂, CO, H₂O, H₂, O₂, ethane, propane, etc), differentoperating pressures, and/or different space velocities may change theselectivity and conversion numbers but that, while the exact level ofimprovement demonstrated by this disclosure may change, the directionalimprovements obtained by this disclosure will still be achieved. Inaddition, it is to be appreciated that, as a basis for the modelingcalculations discussed below, it is assumed that the methane feed to thereactor was always preheated to the same temperature (600° C.) and inall cases a nominal feed rate of methane of 100 kilograms per hour wasused. It was also used as a basis that the catalyst supplied to themoving bed reactor systems was maintained at the same temperature (850°C.). The quantity of catalyst required to maintain this temperature wascalculated for each reactor configuration. For simplicity, it is assumedthat the catalyst thermal conductivity, thermal diffusivity and surfaceemissivity remain constant. The following Table 2 lists the physicalconstants and catalyst properties used in the calculations.

TABLE 2 Model Parameters Catalyst Particle Density 1400 kg/m³ CatalystHeat Capacity 1262 J/kg-K Catalyst Thermal Conductivity 0.4 W/m-KCatalyst Thermal Diffusivity 2.26 × 10⁻⁷ m²/s Catalyst SurfaceEmissivity 0.85

To allow modeling of various reactor configurations, equations wereobtained for Sel_(BTN), Sel_(Coke), and Conv_(BL) by obtaining best fitpolynomial equations for the above set of data points; the data pointswhere H₂ was included in the feed were not included in the calculationsof the equations. The equations obtained and the R² values are shownbelow:Sel_(BTN)=(1.81818181818345E−10)T ⁴−(5.41010101010501E−07)T³+(5.88000000000377E−04)T²−(2.785914141415750E−01)T+4.97583333333585E+01R ² _(BTN)=9.99810335105254E−01Sel_(Coke)=(−1.85878787878687E−10)T ⁴+(5.62280808080511E−07)T³−(6.21721666666349E−04)T ²+(2.996640274168830E−1)T−5.33408809523590E+01R ² _(Coke)=9.99958406639717E−01Conv_(BL)=(1.91428571428569E−06)T²−(1.81714285714283E−03)T+4.53357142857135E−01R ² _(BL)=9.99955208049633E−01

where T is temperature in degrees C.,

In all examples R² is the coefficient of determination which comparesestimated and actual y-values, and ranges in value from 0 to 1. If it is1, there is a perfect correlation in the sample—there is no differencebetween the estimated y-value and the actual y-value. At the otherextreme, if the coefficient of determination is 0, the regressionequation is not helpful in predicting a y-value. The version used hereis based on an analysis of variance decomposition as follows:

$R^{2} = {\frac{{SS}_{R}}{{SS}_{T}} = {1 - {\frac{{SS}_{E}}{{SS}_{T}}.}}}$

In the above definition,

${{SS}_{T} = {\sum\limits_{i}\;\left( {y_{i} - \overset{\_}{y}} \right)^{2}}},{{SS}_{R} = {\sum\limits_{i}\left( {{\hat{y}}_{i} - \overset{\_}{y}} \right)^{2}}},{{SS}_{E} = {\sum\limits_{i}{\left( {y_{i} - {\hat{y}i}} \right)^{2}.}}}$

-   -   That is, SS_(T) is the total sum of squares, SS_(R) is the        regression sum of squares, and SS_(E) is the sum of squared        errors.    -   R² _(BTN) is coefficient of determination for the Sel_(BTN)        correlation,    -   R² _(Coke) is coefficient of determination for the Sel_(Coke)        correlation, and    -   R² _(BL) is coefficient of determination for the Conv_(BL)        correlation.

These set of equations was used to calculate the yields that would beobtained for various reactor configurations where Yield_(BTN) wasdefined as Sel_(BTN)×Conv_(BL) integrated over the temperature profilein the reactor system and Yield_(Coke) was defined asSel_(Coke)×Conv_(BL) integrated over the temperature profile in thereactor system. While it is recognized and shown in the Table 1, thatthe byproduct H₂ improved the reaction selectivity, these equationsomitted the selectivity improvement so that they provided a conservativeestimate as to the level of improvement that the present process wouldprovide.

Transport or Riser Reactor (Comparative)

Utilizing the above equations for a transport or riser reactor withadiabatic declining temperature with an inlet temperature of 850° C. therequired catalyst circulation rate to maintain an outlet temperature of800° C. was 3211 kilograms per hour (kg/hr) based on the nominal feedrate of methane of 100 kg/hr at 600° C. The following yields andselectivities were calculated:Sel_(BTN)=51%Sel_(Coke)=40%Yield_(BTN)=12%Yield_(Coke)=8.9%ΔT_(Reaction)=−50° C. (negative 50° C.);wherein ΔT_(Reaction) is defined as the product outlet reactiontemperature (i.e., the last temperature at which catalytic reactionoccurs before the hydrocarbon product leaves the reactor system) minusthe hydrocarbon feed inlet reaction temperature (i.e., the firsttemperature at which catalytic reaction occurs when the hydrocarbon feedenters the reactor system).Fixed Bed Reactor (Comparative)

Performing modeling of the potential fixed bed comparatives resulted ineven poorer performance than with the transport or riser reactor becausein the fixed bed configuration the entire heat of reaction had to besupplied by the methane containing stream (since no moving catalyst wasused to supply heat to the reaction zone). Therefore the fixed bedreactor required that the methane containing stream had to be heated toa temperature much greater than the desired outlet temperature of 800°C., thereby resulting in a larger magnitude ΔT_(Reaction); that is aΔT_(Reaction) of −60° C. or more negative.

Settling Bed Reactor

In the case simulated for a settling bed of catalyst with an inversetemperature profile and a 50° C. approach temperature between thesupplied catalyst and the process outlet temperature, the inlet wasoperated at 620° C. and the outlet was operated at 800° C., the catalystcirculation rate was reduced to 717 kg/hr and the reaction results wereimproved:Sel_(BTN)=89%Sel_(Coke)=7%Yield_(BTN)=20%Yield_(Coke)=1.5%ΔT_(Reaction)=+180° C.

EXAMPLE 2

Based on the model predicted advantages for an inverse temperatureprofile, a laboratory scale unit was constructed to validate the modelresults. While the model was oriented toward operation of the reactionsystem as a settling bed, the laboratory reactor was a fixed bed ofcatalyst with an inverse temperature profile imposed by use of externalheaters. In all cases the experimentally observed conversions fell belowthe model predicted conversions. This may be due to laboratory scaleexperimental artifacts such as bed bypassing and or/back mixing due tothe hydrodynamic regime in which the lab scale reactors operate.

Mo/ZSM-5 catalyst was prepared via ball milling of 7.5 wt % Mo (wt % ofmetal based on the total weight of the catalyst) as MoO₃ with NH₄ZSM-5support (having a Si/Al₂ ratio of 25) for 2 hr, followed by calcinationat 500° C. for 5 hr in air. The catalyst was pelletized, crushed, andsieved to 20-40 mesh particle size. Catalytic testing of the Mo/ZSM-5catalyst was performed in a fixed bed quartz reactor with an innerdiameter of 7 mm and a bed length of about 18 cm. Inert quartz particles(20-50 mesh) were used as a bed diluent so that all beds were the samelength.

Catalyst performance for methane dehydrocyclization to benzene wasperformed using a 95 vol % CH₄/5 vol % Ar feed (argon was used as aninternal standard). All experimental reaction data was obtained at 20psia (138 kPa-a). The reaction effluent was analyzed using a massspectrometer to determine product concentrations.

Ten separate catalyst performance experiments were conducted forcomparison. In all experiments the catalyst was activated by heating in15 vol % CH₄/80 vol % H₂/5 vol % Ar at 5° C./min to 800° C. and holdingfor 30 min. This was followed by aging the catalyst with 5 cycles ofreaction and regeneration (also identical for all ten experiments). Eachreaction segment lasted 20 minutes at 800° C. in 95 vol % CH₄/5 vol % Arfeed at 1.4 hr⁻¹ weight-hourly space velocity (WHSV) based on CH₄. Eachregeneration segment consisted of switching to H₂, heating to 850° C.with a 10 min. hold time, then cooling back to 800° C. (total time on H₂of 14 min.). The ten experiments differed only on their sixth reactioncycle which was run in 95 vol % CH₄/5 vol % Ar feed for 4 hours.Conditions for the sixth cycle were selected to compare the effects ofcatalyst bed temperature profile at different space velocities. Inparticular, experiments 1 to 5 were run at WHSV values varying between0.25 and 8 hr⁻¹ with bed being held at isothermal conditions at 800° C.In contrast, experiments 6 to 10 were run over the same range of WHSVvalues but with a linear gradient in bed temperature of 650° C. at thefeed inlet to 800° C. at the product outlet (inverse temperatureprofile). Table 3 summarizes the catalyst performance results for theten experiments during reaction cycle #6.

TABLE 3 Increase in WHSV Methane Benzene Benzene Methane Benzene BenzeneTotal Benzene Total (hr⁻¹) for conversion yield (%) selectivityconversion yield (%) selectivity Produced (g Benzene Exp. # cycle 6 (%)at 1 hr at 1 hr (%) at 1 hr (%) at 4 hr at 4 hr (%) at 4 hr C₆H₆/gCatalyst) Produced 1 0.25 17.4 10.3 60 12.5 4.5 36 0.06 Base 2 0.5 17.912.1 68 5.6 0.6 11 0.15 Base 3 1 17.4 11.9 68 1.1 0.0 0 0.24 Base 4 215.1 9.6 64 0.0 0.0 0 0.36 Base 5 8 2.0 0.4 20 0.0 0.0 0 0.38 Base 60.25 16.2 10.0 62 13.2 8.4 64 0.09 50% 7 0.5 16.5 11.0 67 12.6 8.3 650.20 33% 8 1 15.5 11.0 71 9.8 6.4 65 0.35 46% 9 2 12.2 8.2 67 3.8 2.3 600.44 33% 10 8 3.9 2.6 67 0.5 0.4 67 0.72 90%

The results in Table 3 show that there was a clear advantage foroperating with an inverse temperature profile which improvedinstantaneous selectivity at most space velocities for shorter operationtimes and consistently acted to prolong selectivity to benzene overlonger operation times. This allowed for greater cumulative productionin comparison to an isothermal bed at all space velocities

While the above examples are directed to settling bed reactors, similarimprovement in selectivity would be exhibited for other reactor systemswith an equivalent inverse temperature profile.

As demonstrated in the simulation, the yields, selectivities, andcatalyst circulation rates were improved. In addition, the entirereaction may be accomplished in a single reaction zone therebyminimizing required equipment. Optionally, two or more reaction zonesmay be used.

As illustrated by the examples; the settling bed configuration, or otherreactor systems with a similar inverse temperature profile, enables theconversion of methane to higher hydrocarbons, e.g., aromatic compounds,at reduced aging/mechanical-attrition catalyst losses, improvedoperability, and higher selectivity; i.e., lower coke make; than thefixed bed, and/or transport or riser configurations.

Other advantages of the process described and exemplified herein are:

(a) extending catalyst life by minimizing catalyst exposure to hightemperature for a given process outlet temperature due to the inversedtemperature profile;

(b) providing operating flexibilities, such as more reactive species(C₂+) in feed, without substantially increasing in coke formationbecause of the inversed temperature profile, which allows the morereactive feed to contact the cooler/coked catalyst first and convert atlower temperature to desired products rather than more predominatelyconverting to coke at higher temperatures;

(c) enabling the hottest catalyst to be maintained in a more hydrogenrich environment, which reduces coking rate on the hot and/or freshlyregenerated catalyst;

(d) minimizing feed preheat requirements and coking on heat transfersurfaces by direct contacting the with the catalyst;

(e) mitigating reactor metallurgy issues;

(f) adding catalyst continuously to offset catalyst aging;

(g) improving energy efficiency (like a counter current heat exchanger)by reducing catalyst circulation requirements thereby reducing size ofassociated hardware and catalyst attrition; and/or

(h) minimizing product entertainment with exiting catalyst.

While the illustrative embodiments of this disclosure have beendescribed with particularity, it will be understood that various othermodifications will be apparent to and can be readily made by thoseskilled in the art without departing from the spirit and scope of thisdisclosure. Accordingly, it is not intended that the scope of the claimsappended hereto be limited to the examples and descriptions set forthherein but rather that the claims be construed as encompassing all thefeatures of patentable novelty which reside in the present disclosure,including all features which would be treated as equivalents thereof bythose skilled in the art to which this disclosure pertains.

1. A process for converting methane to higher hydrocarbon(s) includingaromatic hydrocarbon(s) in a reaction zone, the process comprising: (a)providing to said reaction zone a hydrocarbon feedstock containingmethane; (b) providing to said reaction zone a catalytic particulatematerial; (c) contacting said catalytic particulate material and saidhydrocarbon feedstock in a substantially countercurrent fashion; (d)maintaining the hydrodynamics of said reaction zone in settling bedregime; and (e) operating said reaction zone under reaction conditionssufficient to convert at least a portion of said methane to a firsteffluent having said higher hydrocarbon(s) wherein said reaction zone isoperated at a gas velocity of less than 0.99 times of the minimumfluidizing velocity.
 2. The process recited in claim 1, furthercomprising steps of separating unreacted methane from said higherhydrocarbon(s) and recycling said unreacted methane to said reactionzone.
 3. The process recited in claim 1, wherein said first effluentalso comprises hydrogen and the process further comprises (i) separatingat least part of said hydrogen from said first effluent or (ii) reactingat least part of said hydrogen from said first effluent withoxygen-containing specie(s) to produce a second effluent having areduced hydrogen content compared with said first effluent.
 4. Theprocess recited in claim 3 further comprising recycling said secondeffluent to (a).
 5. The process recited in claim 1, wherein (a) furthercomprises providing a non-catalytic particulate material to saidreaction zone.
 6. The process recited in claim 5, wherein the mass ratioof the total flowrate of said particulate material (catalyticparticulate material plus any non-catalytic particulate material) to theflowrate of said hydrocarbon feedstock is from about 1:1 to about 40:1.7. The process recited in claim 1, wherein said reaction zone is locatedin a cold wall reactor.
 8. The process recited in claim 1, wherein saidreaction conditions are non-oxidizing conditions.
 9. The process recitedin claim 1, wherein said reaction conditions include a temperature ofabout 400° C. to about 1200° C., a pressure of about 1 kPa-a to about1000 kPa-a, and a weight hourly space velocity of about 0.01 hr⁻¹ toabout 1000 hr⁻¹.
 10. The process recited in claim 1, wherein saidconditions are sufficient to convert at least 5 wt. % of said methane tosaid aromatic hydrocarbon(s).
 11. The process recited in claim 1,wherein said catalytic particulate material comprises a metal orcompound thereof on an inorganic support.
 12. The process recited inclaim 1, wherein said catalytic particulate material comprises at leastone of molybdenum, tungsten, rhenium, a molybdenum compound, a tungstencompound, a zinc compound, and a rhenium compound on ZSM-5, silica or analuminum oxide.
 13. The process recited in claim 1, wherein at least 90wt. % said catalytic particulate material has particle size from about0.1 mm to about 100 mm.
 14. The process recited in claim 1, wherein saidcatalytic particulate material enters said reaction zone at atemperature of about 800° C. to about 1200° C. and exits said reactionzone at a temperature of about 500° C. to about 800° C.
 15. The processrecited in claim 14, wherein the temperature difference of saidcatalytic particulate material across said reaction zone is at least100° C.
 16. The process recited in claim 1, wherein said hydrocarbonfeedstock further comprises at least one of CO₂, CO, H₂, H₂O, or C₂+hydrocarbon(s).
 17. The process recited in claim 1, and furthercomprising: (f) removing at least a portion of said catalyticparticulate material from said reaction zone; (g) regenerating at leasta portion of the removed catalytic particulate material underregenerating conditions; and (h) recycling at least a portion of theregenerated catalytic particulate material to said reaction zone. 18.The process recited in claim 17, wherein said regenerating is conductedat a temperature from about 400° C. to about 750° C. in the presence ofa regeneration gas comprising oxygen.
 19. The process recited in claim18, wherein said regeneration gas further contains carbon dioxide and/ornitrogen such that the oxygen concentration of said regeneration gas isfrom about 2 wt. % to about 10 wt. %.
 20. The process recited in claim17, wherein said regenerating is conducted at a temperature from about800° C. to about 1200° C. in the presence of a regeneration gascomprising hydrogen.